Process for the preparation of propane-1,3-diol by vapor phase hydrogenation of 3-hydroxypropanal, beta-propiolactone, oligomers of beta-propiolactone, esters of 3-hydroxypropanoic acid or mixtures thereof

ABSTRACT

A process is described for the production of propane-1,3-diol. The process comprises subjecting a vaporous feed mixture comprising a hydrogen-containing gas and a feedstock selected from 3-hydroxypropanal, β-propiolactone, oligomers of β-propiolactone, esters of 3-hydroxypropanoic acid, and mixtures of two or more thereof to hydrogenation conditions in a hydrogenation zone in the presence of a heterogeneous hydrogenation catalyst, and recovering a reaction product comprising propane-1,3-diol.

[0001] This invention relates to the production of propane-1,3-diol.

[0002] Propane-1,3-diol is used as an intermediate in the production ofpolyesters for production of fibres or films. It can be prepared by atwo-step process in which ethylene oxide is subjected to an oxonationreaction followed by hydrogenation:

 HOCH₂CH₂CHO+H₂→HOCH₂CH₂CH₂OH

[0003] U.S. Pat. No. 5,981,808 describes the use of anon-phosphine-ligated cobalt compound as oxonation catalyst in anessentially non-water-miscible solvent followed by water extraction toseparate the catalyst from the 3-hydroxypropanal produced as oxonationproduct. The aqueous mixture containing the 3-hydroxypropanal is thensubjected to hydrogenation. U.S. Pat. No. 5,585,528 proposes addition ofa lipophilic tertiary amine as a promoter in such a process. Use ofmethyl t-butyl ether for extraction of the aqueous mixture to recovercobalt catalyst for re-use is described in U.S. Pat. No. 5,770,776. U.S.Pat. No. 5,786,524 teaches a similar process and proposes the use of arhodium catalyst as an alternative catalyst in the oxonation step.

[0004] It is, however, a drawback of such a process that high levels ofbyproducts are produced during liquid phase hydrogenation of theintermediate 3-hydroxypropanal under the recommended hydrogenationconditions, namely liquid phase hydrogenation at 220° C. and 100 bar(1000 kPa). Under such conditions the conversion of 3-hydroxypropanal isonly about 90% while up to 10% is converted to byproducts.

[0005] It has also been proposed to combine the oxonation andhydrogenation steps into a one-step process with, it is claimed, minimalproduction of 3-hydroxypropanal as byproduct. Such a one-step processcan be effected using a phosphine complex of cobalt carbonyl as themajor catalyst ingredient. However, the use of a ruthenium compound ascatalyst has also been proposed. An organic solvent is used in thereaction enabling a water extraction to be used in order to separatepropane-1,3-diol from the oxonation mixture. Ethylene oxide conversionsof 55% with a selectivity towards propane-1,3-diol of 87% are reported.

[0006] U.S. Pat. Nos. 5,310,948 and 5,359,081 teach formation ofβ-propiolactone or polymers thereof by reaction of carbon monoxide andethylene oxide in the presence of a cobalt-containing catalyst systemcomprising a source of cobalt and a hydroxyl-substituted pyridine.

[0007] Propane-1,3-diol can alternatively be produced from glycerolusing recombinant bacteria expressing recombinant diol dehydratase. Sucha process is taught in U.S. Pat. No. 5,821,092.

[0008] It has also been proposed to subject acrolein to hydration so asto form 3-hydroxypropanal which is then hydrogenated. In this connectionreference may be made to U.S. Pat. No. 5,364,987.

[0009] In U.S. Pat. No. 5,334,778 it is proposed to producepropan-1,3-diol having a residual carbonyl content below 500 ppm bycatalytically hydrogenating 3-hydroxypropanal in aqueous solution in thepresence of a hydrogenation catalyst at 30° C. to 80° C. to a3-hydroxypropanal conversion of 50% to 95% and then continuing thehydrogenation at 120° C. to 140° C. to achieve a 3-hydroxypropanalconversion of substantially 100%.

[0010] Both glycerol and acrolein are, however, generally more expensiveand less available than ethylene oxide. Hence it is not currently aneconomic proposition to manufacture propane-1,3-diol by either of theselast two mentioned processes.

[0011] It would be desirable to provide an improved process for theproduction of propane-1,3-diol. It would further be desirable to providea process for producing propane-1,3-diol by hydrogenation of anappropriate carbonyl compound which exhibits increased selectivitytowards propane-1,3-diol and reduced amounts of undesirable byproducts,such as propan-1-ol, which cannot readily be converted topropane-1,3-diol. It would further be desirable to provide a process forthe production of propane-1,3-diol by hydrogenation of an intermediatecompound which can be made from ethylene oxide and which contains atleast one carbon-oxygen double bond, such as β-propiolactone, oligomersof β-propiolactone, or an ester of 3-hydroxypropionic acid, with minimalformation of undesirable byproducts.

[0012] It is accordingly an objective of the present invention toprovide an improved process for the production of propane-1,3-diol. Inaddition the present invention seeks to provide a hydrogenation processfor producing propane-1,3-diol that uses in a hydrogenation step anoptimised catalyst system. Yet a further objective of the presentinvention is to provide a process for the production of propane-1,3-diolby hydrogenation of an intermediate compound which can be made fromethylene oxide, such intermediate compound containing at least onecarbon-oxygen double bond, and preferably being selected from,β-propiolactone, oligomers of β-propiolactone, esters of3-hydroxypropionic acid, and mixtures of two or more thereof, withreduced amounts being formed of undesirable byproducts, such aspropan-1-ol, which cannot readily be converted to the desiredpropane-1,3-diol.

[0013] According to the present invention there is provided process forthe production of propane-1,3-diol which comprises forming a vaporousfeed mixture comprising a hydrogen-containing gas and a substantiallyanhydrous feedstock selected from β-propiolactone, oligomers ofβ-propiolactone, esters of 3-hydroxypropanoic acid, and mixtures of twoor more thereof, supplying the vaporous feed mixture to a hydrogenationzone containing a heterogeneous hydrogenation catalyst selected fromreduced copper oxide/zinc oxide hydrogenation catalysts, reducedmanganese-promoted copper catalysts, reduced copper chromite catalystsand reduced promoted copper chromite catalysts at a temperature of fromabout 130° C. to about 10° C. and a feed pressure to the hydrogenationzone of from about 50 psia (about 344.74 kpa)to about 2000 psia(13789.52 kPa) said hydrogenation conditions effective for hydrogenatingfeedstock to propane-1,3-diol, and recovering from the hydrogenationzone a reaction product comprising propane-1,3-diol.

[0014] The feedstock to the hydrogenation step is substantiallyanhydrous, that is to say it contains no more than about 5% (w/v),preferably no more than about 1% (w/v), and even more preferably lessthan about 0.1% (w/v) of water. It is selected from β-propiolactone,oligomers of β-propiolactone, and esters of 3-hydroxypropanoic acid, andmixtures of two or more thereof. β-propiolactone can self-polymerise toform oligomers of β-propiolactone. The presence of more than a minoramount of such oligomers in the feedstock to the hydrogenation zone isgenerally undesirable because of their relative lack of volatility.Hence it will normally be preferred to use a feedstock to thehydrogenation zone which contains less than about 10 molar % ofoligomers of β-propiolactone. Accordingly it will usually be preferredto minimise the proportion of oligomers of β-propiolactone in thefeedstock to the hydrogenation zone.

[0015] Hydrogenation is effected using a vaporous feed mixture to thehydrogenation zone, this mixture containing in addition to the feedstockalso a hydrogen-containing gas. The hydrogen-containing gas ispreferably substantially free from carbon oxides but may contain one ormore inert gases, such as nitrogen, argon and helium, in amounts of upto 50% v/v, but which preferably do not exceed about 10% v/v and morepreferably do not exceed about 5% v/v, e.g. about 1% v/v or less.

[0016] The hydrogenation conditions may be selected so that the reactionmixture exiting the hydrogenation zone is also in the vapour phase.However, it is alternatively possible, and indeed may be preferable, toutilise hydrogenation conditions such that the reaction mixture at theexit end of the hydrogenation zone is below its dew point so that atleast some of the condensable components thereof are present in theliquid phase.

[0017] The hydrogenation zone conveniently contains a fixed bed of agranular hydrogenation catalyst. If desired, the hydrogenation zone cancontain more than one catalyst bed and the hydrogenation catalyst of onebed can, if desired, differ from the hydrogenation catalyst of at leastone other bed.

[0018] The active catalytic species in the hydrogenation catalyst may beat least partially supported on a supporting material selected fromchromia, zinc oxide, alumina, silica, silica-alumina, silicon carbide,zirconia, titania, carbon, or a mixture of two or more thereof, forexample, a mixture of chromia and carbon.

[0019] Preferably the hydrogenation catalyst is a reducedmanganese-promoted copper catalyst. Such manganese-promoted coppercatalysts preferably have a total surface area of at least about 15m²/g, more preferably at least about 20 m²/g, and even more preferablyat least about 25 m²/g, in the unreduced form.

[0020] A particularly preferred hydrogenation catalyst is a reducedmanganese-promoted copper catalyst which is available as DRD92/89Acatalyst from Kvaerner Process Technology Limited of The TechnologyCentre, Princeton Drive, Thornaby, Stockton-on-Tees, TS17 6PY, England.Alternatively there may be used a reduced manganese-promoted coppercatalyst which is available as DRD92/89B catalyst from Kvaerner ProcessTechnology Limited.

[0021] The hydrogenation step is conducted under vapour phase feedconditions so that the feed stream to the hydrogenation zone is aboveits dew point and is thus a vaporous feed stream. The reaction productmixture from the hydrogenation zone can also be recovered at atemperature above its dew point so that it too is in vaporous form or itcan be recovered at a temperature below its dew point so that at leastpart of the condensable components thereof are in liquid form.

[0022] Although it is possible to conduct the hydrogenation process ofthe present invention in a tubular reactor under substantiallyisothermal conditions, it will normally be preferred to operate undersubstantially adiabatic hydrogenation conditions using a fixed catalystbed or beds since adiabatic reactors are much cheaper to construct thantubular reactors. However, care should be taken in designing the plant,in particular in selecting a suitable gas:feedstock ratio and inchoosing a catalyst size, that the temperature rise across any catalystbed is limited to a reasonable value, typically not more than about 20°C., so as to keep the temperature to which the reaction mixture isexposed within desired limits. In this way the formation of 1-propanolas a byproduct can be limited.

[0023] It will normally be preferred that, in the vaporous feed streamto the hydrogenation zone, the hydrogen-containing gas:feedstock molarratio shall be in the range of from about 50:1 to about 1000:1.

[0024] Typically the feed temperature to the hydrogenation zone is fromabout 130° C. to about 180° C., more preferably from about 135° C. toabout 150° C., while the feed pressure to the hydrogenation zone is fromabout 50 psia (about 344.74 kPa) to about 2000 psia (about 13789.52kPa), for example, from about 350 psia (about 2413.17 kPa) to about 1000psia (about 6894.76 kPa). The feedstock is also preferably supplied tothe first hydrogenation zone at a rate corresponding to a liquid hourlyspace velocity of from about 0.05 to about 5.0 h⁻¹. Preferably unreactedhydrogen-containing gas is recycled for further use.

[0025] If desired, the feedstock to the hydrogenation zone can bediluted with a solvent, such as methanol, which is stable under thehydrogenation conditions utilised.

[0026] The invention is further illustrated in the following Examples.

EXAMPLE 1

[0027] This Example was intended to provide a simulation of conditionssimilar to those that might exist in a commercial hydrogenation reactorfor hydrogenation of methyl 3-hydroxypropionate utilising recycledhydrogen which would be saturated with methanol co-product.

[0028] A solution of crude methyl 3-hydroxypropionate (approximately 90%pure) was diluted with an approximately equal weight of methanol to forma feed solution. This feed solution was subjected to vapour phasehydrogenation in a once-through adiabatic fixed-bed reactor. The reactorwas constructed from a 95 cm length of 20.96 mm internal diameter tubewhich was oil-jacketed to reduce heat losses. The reactor contained acharge of 100 ml of DRD92/89A catalyst which is obtainable from KvaernerProcess Technology Limited of The Technology Centre, Princeton Drive,Thornaby, Stockton-on-Tees, TS17 6PY, England. The catalyst was reducedby a procedure analogous to that described in U.S. Pat. No. 5,030,609.

[0029] The feed solution was supplied at a feed rate of 12 ml/h to aheater and vaporised by a stream of pure hydrogen at a rate of 1000 Nl/h(i.e. 1000 litres per hour measured at 760 mm Hg [101.33 kPa] and 0°C.). The vaporous mixture was passed over the catalyst at a pressure of400 psig (2757.90 kPa gauge) and a temperature of 138° C. The reactionproduct mixture exiting the reactor was cooled and the condensed liquidcollected. The feed and product were analysed by gas chromatographyusing a 60 metre CP SIL 19 capillary column of 0.32 mm internal diameterwith a 1.3 μm film thickness.

[0030] Conversion of the methyl 3-hydroxypropionate was determined as73.3% with a selectivity to propane-1,3-diol of 83.4% and to 1-propanolof 5.8%. It is believed that the byproducts comprise mainly materialswhich, upon hydrogenation, can be converted to propane-1,3-diol andhence can be recycled.

EXAMPLE 2

[0031] The general procedure of Example 1 was repeated except that thetemperature was 148° C. Conversion of the methyl 3-hydroxypropionate wasdetermined as 84.9% with a selectivity to propane-1,3-diol of 84.8% andto 1-propanol of 7.2%. Operation at this slightly higher temperatureprovides an expected increase in conversion with a minor increase inselectivity to 1-propanol.

EXAMPLE 3

[0032] The general procedure of Example 1 was repeated except that thetemperature was 174° C. Conversion of the methyl 3-hydroxypropionate wasdetermined as 99.95% with a selectivity to propane-1,3-diol of 13.1% andto 1-propanol of 81.8%. This shows that operation at high temperaturefavours formation of the alcohol, 1-propanol, rather thanpropan-1,3-diol.

EXAMPLE 4

[0033] The same feed solution as used in Example 1 was subjected tovapour phase hydrogenation in an adiabatic fixed-bed reactor systemwhich incorporated recycle of excess gas following condensation of thereactor product stream. Make-up pure hydrogen was supplied to the systemto maintain constant system pressure. The reactor was constructed from a200 cm length of 26.64 mm internal diameter tube which was insulated andprovided with electric trace heating means to prevent net heat loss fromthe reactor. The reactor contained a charge of 250 ml of DRD92/89Acatalyst. The catalyst was reduced by a procedure analogous to thatdescribed in U.S. Pat. No. 5,030,609.

[0034] The feed solution was fed at a rate of 80 ml/h to a heater andvaporised by a stream of mixed recycle gas and pure hydrogen at a rateof 10900 Nl/h. The vaporous mixture was passed over the catalyst at apressure of 885 psig (6101.86 kPa gauge) and an inlet temperature of149° C. The outlet temperature was measured as 150° C. The reactorproduct mixture was cooled and the condensed liquid collected. The feedand product were analysed by gas chromatography using a 60 metre CP SIL19 capillary column of 0.32 mm internal diameter with a 1.3 μm filmthickness.

[0035] Conversion of the methyl 3-hydroxypropionate was determined as77.8% with a selectivity to propane-1,3-diol of 78.9% and to 1-propanolof 14.8%.

EXAMPLE 5

[0036] The general procedure of Example 4 was repeated except that theinlet pressure was 735 psig (5067.65 kPa), and the mixed recycle andpure hydrogen make-up flow rate was 8116 Nl/h. Conversion of the methyl3-hydroxypropionate was determined as 61.1% with a selectivity topropane-1,3-diol of 79.0% and to 1-propanol of 14.4%. Operation at thislower pressure and lower gas rate provides an expected reducedconversion but with similar selectivities to propane-1,3-diol and to1-propanol.

EXAMPLE 6

[0037] A solution of crude methyl 3-hydroxypropionate (approximately 98%pure) was diluted with an approximately equal weight of methanol to forma feed solution. This feed solution was subjected to vapour phasehydrogenation in the apparatus used in Example 4.

[0038] The feed solution was fed at a rate of 80.8 ml/h to a heater andvaporised by a stream of mixed recycle gas and pure hydrogen at a rateof 8717 Nl/h. The vaporous mixture was passed over the catalyst at apressure of 885 psig (6101.86 kPa gauge) and an inlet temperature of148° C. The outlet temperature was measured as 149° C.

[0039] The reactor product mixture was cooled and the condensed liquidcollected.

[0040] The feed and product were analysed by the method used in Example4.

[0041] Conversion of the methyl 3-hydroxypropionate was determined as71.3% with a selectivity to propane-1,3-diol of 80.6% and to 1-propanolof 12.5%. It is believed that the byproducts comprise mainly materialswhich, upon hydrogenation, can be converted to propane-1,3-diol andhence can be recycled.

EXAMPLE 7

[0042] The general procedure of Example 6 was repeated except that thefeed rate was 61.4 ml/h, the recycle stream was 6537 Nl/h and the inlettemperature was 149° C. The outlet temperature was measured as 150° C.

[0043] Conversion of the methyl 3-hydroxypropionate was determined as84.1% with a selectivity to propane-1,3-diol of 73.7% and to 1-propanolof 21.5%.

EXAMPLE 8

[0044] The general procedure of Example 7 was repeated except that theinlet temperature was 154° C. The outlet temperature was measured as155° C.

[0045] Conversion of the methyl 3-hydroxypropionate was determined as87.3% with a selectivity to propane-1,3-diol of 75.2% and to 1-propanolof 18.7%.

EXAMPLE 9

[0046] The general procedure of Example 8 was repeated except that themixed recycle and pure hydrogen make-up flow rate was 5226 Nl/h and theinlet temperature was 152° C. The outlet temperature was measured as153° C.

[0047] Conversion of the methyl 3-hydroxypropionate was determined as86.7% with a selectivity to propane-1,3-diol of 76.2% and to 1-propanolof 16.9%.

1. A process for the production of propane-1,3-diol which comprisesforming a vaporous feed mixture comprising a hydrogen-containing gas anda substantially anhydrous feedstock selected from β-propiolactone,oligomers of β-propiolactone, esters of 3-hydroxypropanoic acid, andmixtures of two or more thereof, supplying the vaporous feed mixture toa hydrogenation zone containing a heterogeneous hydrogenation catalystselected from reduced copper oxide/zinc oxide hydrogenation catalysts,reduced manganese-promoted copper catalysts, reduced copper chromitecatalysts and reduced promoted copper chromite catalysts at atemperature of from about 130° C. to about 180° C. and a feed pressureto the hydrogenation zone of from about 50 psia (about 344.74 kpa)toabout 2000 psia (13789.52 kPa) said hydrogenation conditions effectivefor hydrogenating feedstock to propane-1,3-diol, and recovering from thehydrogenation zone a reaction product comprising propane-1,3-diol.
 2. Aprocess according to claim 1 wherein the feedstock comprises an alkylester or a hydroxyalkyl ester of 3-hydroxypropanoic acid.
 3. A processaccording to claim 1 or claim 2, wherein hydrogenation is effected usinga fixed bed of granular hydrogenation catalyst.
 4. A process accordingto any one of claims 1 to 3, wherein the hydrogen-containinggas:feedstock molar ratio in the vaporous feed mixture is in the rangeof from about 50:1 to about 1000:1.
 5. A process according to any one ofclaims 1 to 4, wherein the feed temperature to the hydrogenation zone isfrom about 135° C. to about 150° C.
 6. A process according to any one ofclaims 1 to 5, wherein the feed pressure to the hydrogenation zone isfrom about 350 psia (about 2413.17 kPa) to about 1000 psia (about6894.76 kPa).
 7. A process according to any one of claims 1 to 6,wherein the feedstock is supplied to the first hydrogenation zone at arate corresponding to a liquid hourly space velocity of from about 0.05to about 5.0 h⁻¹.